Benzene production



Dec. 30, 1969 R. G. CRAIG ET AL 3,487,120

BENZENE PRODUCTION Filed Dec. 16. 966 2 Sheets-Sheet 1 IN VENTOR.

Q ROBERT &- creme LOUIS c. 3054. BY LEE FRIEDMAN ATToRNEy Sheets-Sheet INVENTORS. ROBERT G. CRAIG LOUIS O. DOE'LP R. G. CRAIG ET AL BENZENE PRODUCTION LEE FRIEDMAN Dec. 30, 1969 Filed Dec. 16, 1966 Maw ATTOR Y.

United States Patent 3,487,120 BENZENE PRODUCTION Robert G. Craig, Wilmington, Del., and Louis C. Doelp, Jr., Glen Mills, and Lee Friedman, Bala Cynwyd, Pa., assignors to Air Products and Chemicals, Inc., Philadelphia, Pa., a corporation of Delaware Filed Dec. 16, 1966, Ser. No. 602,323 Int. Cl. C07c 7/02, 3/58 US. Cl. 260674 9 Claims ABSTRACT OF THE DISCLOSURE This invention relates to a process for the treatment of pyrolysis gasoline and Similar olefinic charge stocks and, more specifically, to the catalytic hydrogenation, hydrocracking and dealkylation of pyrolysis condensates, such as a dripolene, particularly for the production of benzene and concomitant fuel gas.

Dripolene is the normally liquid mixture of hydrocarbons obtained as a byproduct in the production of ethylene and/or propylene during the high temperature pyrolysis of gaseous and/0r liquid hydrocarbons. A material which condenses out of the hydrocarbons when the pyrolysis products are rapidly cooled, e.g., by quenching, is a liquid known as dripolene, pyrolysis condensate or pyrolysis naphtha, which generally boils between 7.5 to 500 F. The expression pyrolysis gasoline is also employed generically to define compositions which include dripolene.

While suggestions have been made to send dripolene to aromatic extraction units for recovery of pure aromatics such as benzene, without further treatment the material is unsuitable for use in aromatics production due to its high olefin content including acyclic and cyclic diolefins. Even the potential of dripolene as a blending stock for gasolines has not been realizable. Since dripolene contains materials such as olefins and diolefins which tend to form gum-like polymers, especially when heated, the quantity of dripolene blended into motor fuels has been relatively small.

Various suggestions directed to the selective hydrogenation of mono-olefins and/or diolefins, without simultaneously hydrogenating the aromatic compounds present in dripolene, have not been commercially satisfactory due to such problems as excessive polymer formation and catalyst deactivation.

SUMMARY OF THE INVENTION An object of the present invention is to provide a continuous integrated process for treating light dripolene naphthas to obtain high purity benzene.

It is a further object of the invention to provide a catalytic method for the selective hydrogenation, hydrocracking and dealkylation of pyrolysis condensates or naphthas to provide benzene and fuel gas.

In accordance with the present invention, the catalytic treatment of pyrolysis gasoline charge stock, generally boiling between about 100 and about 350 F and prefer- 3,487,120 Patented Dec. 30, 1969 ICC ably between 150 and 300 F., is effected in at least two reaction zones by first contacting a commingled stream of vaporized charge stock and a hydrogen-containing gas with sulfided-cobalt molybdate or sulfided-nickel cobalt molybdate in a first reaction zone under a carefully regulated set of conditions. Substantially complete hydrogenation of olefinic unsaturates occurs in this first reaction zone and reduces the exothermic heat burden in the second reaction zone. The resulting hydrogenated material is then contacted with chrome-alumina catalyst in a second reaction zone under regulated conditions where hydrocracking of non-aromatics and dealkylation of alkyl benzenes occurs. The resulting effluent from the second reaction zone is cooled, stabilized, clay treated and then fractionated to obtain a finished benzene product. The toluene and/or xylenes remaining in the product may be ecovered for use or recycled to dealkylation.

The catalyst employed in the practice of the present rhvention for the first reaction Zone is an alumina-supported cobalt molybdate or nickel cobalt molybdate which is in a sulfided state during its use. The catalyst may either be presulfided or, depending on the nature of the charge stock, permitted to attain its sulfided state during the initial startup of the operation. The catalyst contains a total of about 8 to 20% by weight of the oxides of the molybdenum and cobalt in the approximate weight ratio of about 3.5/1 to 5/1 (MoO /CoO). The preferred catalysts are those containing at all times during use at least one atom of sulfur for each atom of cobalt and for each atom of molybdenum. Thus, for a cobalt molybdate catalyst which, prior to sulfidation, contains 15% M00 and 3% C00, the minimum sulfur content of the freshly sulfided catalyst will be about 7.9% by weight. In the hereinafter described treatment, this catalyst is not subject to deactivation by sulfur or nitrogen in the feed or hydrogen stream and, in fact, contributes to the removal of such contaminants by converting them into readily separable materials. To maintain the cobalt molybdate catalyst in sulfided state here should be present in the reaction zone a quantity of sulfur compound in excess of theroetical equilibrium requirements and generally corresponding to a partial pressure of H 8 of at least of the hydrogen partial pressure. For example, a content of 0.005% to 0.5% sulfur by weight of hydrocarbon in the feed should generally be sufiicient to provide the required partial pressure of hydrogen sulfide in the reactor to maintain the catalyst in the desired state of sulfidation. In processing of charge stocks of low sulfur content, such as those containing insufficient sulfur compounds to maintain the catalyst in sulfided state, it is preferred to add H 8 or other simple sulfur compound, such as CS or a mercaptan, with the charge to maintain the desired concentration of sulfur in the catalytic reaction zone.

The conditions of temperature, pressure, liquid hourly space velocity and hydrogen-containing gas rate are interrelated. The inlet temperature in the first reaction zone is maintained between 400 and 650 F. and preferably between 500 and 600 F. Pressures from about 600 up to 1200 pounds per square inch gauge can be utilized with pressures from 800 to 1000 pounds per square inch gauge being preferred. A liquid hourly space velocity (LHSV) of 0.1 to 5.0 may be employed with a preferred LHSV in the range of 0.5 to 1.5. The hydrogen-containing gas may be derived from a naphtha reforming operation but, in any respect, is composed of at least 25% of volume hydrogen and preferably should contain about or more of hydrogen. The hydrogen is employed in a hydrogen to oil mol ratio of 1 to 12 and preferably in a mol ratio of 5 to 9'.

The temperature of the charge stock in passing through the first reaction zone rises because of the exothermic nature of the reaction to an outlet temperature above that of the inlet temperature but is not allowed to rise in excess of an outlet temperature of about 750 F. If more than one reactor is employed in the first reaction zone, that total effiuent from the first and each intermediate reactor of the first reaction zone is immediately cooled to an appropriate lower temperature before being admitted to the next succeeding reactor wherein the reactants again contact catalyst.

Since in addition to less of valuable product the catalyst is deactivated as a result of polymerization of unsaturates When excessive temperatures are encountered, the inlet temperature of the charge to the hydrogenation reactor, or to each reactor in the system if more than one is employed, is selected to be as low as reasonably possible so as to approach but not to exceed during the ensuing reaction, the predetermined maximum temperature desired at the reactor outlet, considering the expected exothermic temperature elevation produced in the reaction.

Hydrocracking and hydrodealkylation occur in the second reaction zone wherein hydrogenated effluent from the first reaction zone (e.g. one to two reactors) together with the desired amount of recycled material comprising charge stock not converted to benzene in the process, are contacted with chromia-alumina catalyst in at least one reactor (e.g. one to four reactors). The operating conditions are so selected that the total vapor residence time (based on the empty reactors) of the reactants in the reactors of the second reaction zone lies in the approximate range of 30 to 180 seconds. For the particular catalyst employed, reactor outlet temperatures in the range of 1100 to 1200 F. are generally recommended. The inlet temperature to each reactor is so selected that the outlet temperature resulting from the exothermic reaction is maintained below 1225 F. The efiluent from each reactor in the second reaction zone may be cooled by as much as 100 to 200 F. before the efiluent is passed to the next reactor. A hydrogen level is maintained in these reactors such that there is provided an inlet hydrogen to total aromatics mol ratio of 3:1 to :1 and preferably 7:1. The total pressure maybe varied widely, but is maintained above 500 p.s.i.g. and generally in the range of 800 to 1200 p.s.i.g.

The hydrogen partial pressure is at least 250 p.s.i.g. and may range from 250 to 700 p.s.i.g.

The catalysts employed in the second reaction zone are made by dehydration of an alumina hydrate composition containing at least 50% and preferably in excess of 75% beta trihydrate. After adjusting the surface area by heat treatment of the dehydrated material to the range of 80 to 300 square meters per gram, it is impregnated with chromic acid in an amount furnishing to 25% (Jr- 0 by Weight of the finished catalyst, followed by drying and calcining. The calcined chrome alumina catalyst has a surface area of about 50 to 150 square meters per gram.

With continued operation over comparatively long periods, it may be desirable to raise the operating temperature to some extent in the first reaction zone, the second reaction zone, or both. Periodic increases in temperature will primarily be dictated by the required properties of the unit liquid product.

In addition, the size of the several reactors of the series and the quantity of catalyst therein need not be equal. Accordingly, the temperature to which the products from a preceding reactor are cooled before admission to the next reactor will be governed by the extent of the expected temperature rise therein and consistent with achieving the desired product.

Under these conditions, the catalyst is maintained at a high catalytic activity level which is important not only from the standpoint of the immediate benefits in the yield of recovered valuable liquid product but also, and more importantly, from the standpoint of enhanced uninterrupted on-stream periods of operation.

4 BRIEF DESCRIPTION OF THE DRAWING The invention is clarified by reference to the following description read in connection with the drawings. FIG- URES 1 and 2 are highly schematic illustrations of specific embodiments for the treatment of pyrolysis gasolines.

DESCRIPTION OF PREFERRED EMBODIMENTS In FIGURE 1, pyrolysis gasoline is passed to a vaporizer either directly or indirectly. When pyrolysis condensate or pyrolysis naphtha charge stock is employed, valves 17 and 1 8 are closed and pyrolysis condensate from storage area 10 is passed through lines 11 to 14 directly to venturi device 15, which is attached to vaporizer 16.

In a preferred operation, however, a particular fraction may be separated from a full-range pyrolysis gasoline charge stock in a feed-splitter or fractionator and sent to the vaporizer. In this preferred operation, valves 19 and 20 are closed and pyrolysis gasoline from storage area 10 is conveyed by lines 11 and 21 to feed-splitter 22. In the feed-splitter, a particular fraction, such as the fraction boiling between about and about 300 F., is separated from the charge stock and then transmitted to venturi device 15 through lines 23, 13 and 14. The lower boiling fraction from the feed-splitter is composed mainly of C and lighter hydrocarbons and is removed by line 24- for further separation and/or use as fuel. C and heavier bottoms are removed from the feed-splitter (by means not shown) for use primarily as fuel.

In either embodiment, the pyrolysis gasoline from storage area 10 is preferably raised to the desired operating pressure and may be heated to from 200 to a maximum of about 300 F. (by means not shown) prior to its introduction into vaporizer 16. If desired, sulfur can be added, e.g., as carbon disulfide, to maintain a sulfur level in the charge stock in line 13 which will minimize polymerization in vaporizer 16 and keep the catalyst sulfided in reaction zone 31.

It has been found that when the sulfur concentration of the charge passed through the vaporizer is maintained at about 0.01% based on weight of hydrocarbons, the sulfur compound acts as a polymerization inhibitor and prevents deposition of polymer in the vaporizer and in the connecting lines, in addition to maintaining the cobalt molybdate catalyst in sulfided state. While no deleterious effect has been observed from the presence of excess sulfur, no additional advantages is obtained in adding sulfur compounds to bring the total to above about 0.05%. While not limited thereto, the use of the sulfur polymerization inhibition is particularly important in the vaporization of charge stocks containing 10% or more of readily polymerizable materials such as cyclodiolefins and styrene type compounds.

The charge stock supplied by line 13 may be blended with liquid from line 25, which has not vaporized in vaporizer 16. Generally, however, valve 26 is closed and charge stock from line 13 is passed through line 14 to venturi device or other suitable mixing or contacting means 15, where it is mixed with hot (e.g., 700 to 1100 F.) hydrogen-containing recycle gas from line 27. The commingled mixture of charge stock and hydrogen-containing gas is flashed in vaporizer 16 which is operated between the pressure limits of about 800 to 1200 lbs./in. gauge and a temperature between about 350 to 550 P. where about 0.5 to 4 volume percent and generally 1 to 2 volume percent of the liquid remains unvaporized. The excess accumulation of unvaporized charge is drawn off through line 28, regulated by valve 29 and is discarded or recycled to the feed-splitter (by means not shown).

Vapor mixture passes from the vaporizer through line 30 and after any required adjustment in temperature (e.g. by a heater or by admixture with additional hydrogen, not shown) enters the first reaction zone 31 where the mixture contacts sulfided cobalt molybdate catalyst. The eflluent from the first reaction zone is then passed through line 32 into a heater 33 and thereafter transmitted by line 34 to the second reaction zone 35 containing chromealumina catalyst. The operating conditions for the first and second reaction zone have been set forth above.

Reactor eflluent from the second reaction zone is heat exchanged with recycle gas and possibly other sources in heat exchanger 37 and then normally cooled to about 80 F. in a second heat exchanger 39, after passing through lines 36 and 38, respectively. The cold reactor effluent is passed through line 40 to a gas-liquid separator 41. Gas escapes from the separator through line 42 and together with added makeup hydrogen from source 43 (and line 44) moves through line 45 to recycle pump or compressor 46. The recycled hydrogen-containing gas, after passing through line 47 and heat exchanger 37, is then passed through line 48 and heater 49 to attain the required temperature before being sent back to venturi device by line 27. Unless purifying means are provided for the recycled gas stream, a portion of the separated gas should be discharged (for use as fuel) and replaced by fresh hydrogen to maintain desired purity.

Liquid from the gas-liquid separator 41 is sent by line 50 to stabilizer 51. The overhead from the stabilizer ordinarily passes through line 52 to a gas plant for use as fuel, while liquid is conveyed from the bottom of the stabilizer through line 53 to a clay treater and a benzene fractionation tower. The bottoms fraction from the henzene fractionaton tower can be recycled to the vaporizer, the feed-splitter or preferably line 32 by suitable means. The benzene tower bottoms contain diphenyl and the recycle of this diphenyl suppresses additional diphenyl formation and increases selectivity to benzene. If desired, toluene can be recovered as product from the bottoms fraction.

It will be understood that more than one reactor or catalyst bed may be employed for each reaction zone upon making suitable adjustment of conditions as to the extent of conversion desired to be accomplished in each reactor. Greater flexibility and better control of the operation are best obtained by the use of several reactors. Separate reactors may utilize the same housing.

The embodiment shown in FIGURE 2 illustrates a slightly modified alternative process layout and in addition, identifies certain of the equipment not specifically illustrated in FIGURE 1. Like parts follow the same numbering as in FIGURE 1. Thus, the full pyrolysis naptha from line 11 is fractionated in splitter 22 to obtain the de sired fraction (say 150 to 300 F.) used as charge to the process. A C and lighter stream is discharged in line 24 While the heavy bottoms (C are rejected to fuel or for other use through line 200.

The selected side cut is withdrawn from the feed splitter by line 23 and passed to vaporizer 1. Sulfur in the form of CS or mercaptan may be added to this charge through line 201. The hot hydrogen stream from line 27 is sparged directly into the vaporizer 16, and is made up of fresh hydrogen from source 43 and recycle hydrogen from separator 41, both preheated to required temperature in furnace 49. As in the previously described embodiment, the vaporized portion, comprising 96 or more volume percent of the hydrocarbons admitted through line 14, are sent to the hydro-pretreater 31 with intermediate temperature adjustment when necessary or desired. The unvaporized portion may be discharged through line 28 or may be returned to fractionator 22 by line 204.

The hydrogenated efliuent from 31, as before, is subjected to treatment over chrome-alumina catalyst to effect cracking out of non-aromatics and dealkylation of alkyl aromatics. In the embodiment illustrated in FIGURE 2, two reactors 35a and 35b containing chrome-alumina catalyst are shown with intermediate cooling of the eflluent in line 34b by means of a relatively cool recycled gas stream from compressor 46 passed through line 205. While two reactors 35a and 35b are illustrated, it will be understood that 3 or 4 reactors may be employed to accommodate the required hydrocracking capacity.

As before, the effluent from the last chrome-alumina reactor is cooled and the condensed liquid separated for recovery of benzene. In the presently described alternative scheme, the eflluent in line 36 is first cooled by cold-water at 207 providing for steam generation and the thus precooled stream further cooled at 39 to effect condensation of the normally liquid hydrocarbons therein; liquid-gas separation being effected at 41. If desired, all or part of the recycle gas discharged through line 42 may be purified in any desired manner as shown at 208, before recompression.

The liquid condensate is sent by line 50 to a stabilizer 51 wherein light hydrocarbons are removed and a stabilized liquid product obtained which is concentrated in aromatic hydrocarbons. In preferred practice the arcmatics concentrate is treated over clay in 209 and fractionated in tower 210 to obtain a benzene product of high purity sent to storage by line 211. The higher boiling hydrocarbon bottoms discharged in line 212 may be recycled to the dealkylation reactors (35a), or if desired, a portion may be used to quench the stream in line 34b between reactors 35a and 35b.

The process of the invention can be employed to obtain good yields of high purity benzene not only from pyrolysis gasoline but also from other impure hydrocarbon mixtures of suitable boiling range which are rich in olefins and which are otherwise diflicult to handle because of their relatively high content of polymerizable unsaturates such as cyclopentadienes, dicyclopentadienes, styrenes and indenes, which unsaturates may often constitute as high as 10 to 40% of the feed stock. Among such feed stocks are included coker-naphthas as well as coke-oven derived oil fractions whether of high or low sulfur and nitrogen content, and any stocks containing relatively large amounts of non-aromatics.

The invention will be illustrated by the following specific examples, it being understood that there is no intention to be necessarily limited by any details thereof since variations may be made within the scope of the appended claims.

EXAMPLE 1 A dripolene charge stock boiling over the range of 127 to 293 F. had the following characteristics:

Specific gravity, API 44.6 Total nitrogen, p.p.m. 5

dripolene of the characteristics and composition set forth above was contacted with sulfided cobalt molybdate catalyst, containing prior to sulfidation 15% M00 and 3% C00, in a reactor at the following conditions:

Run (a) (b) gemperature, F 550 450 ressure, p.s.i.g 600 Space rate, LHSV 3 2% Hydrogen to oil ratio, mol/mol 4 2 Under the aforementioned operating conditions, the initial charge stock was hydrogenated to a bromine number of 35 for product (a) and 47 for product (1)). The products of this partial hydrogenation of the charge stock after distillation of an approximate volume percent over-head had the characteristics shown in the following table.

TABLE 1 Product Charge (a) (b) Bromine No 64. 5 35 47 After maleic anhydride treat- 39. 4 35 47 Existent gum, mg./l ml. 86. 4 n-Hexane washed 0. 1 0. 6 16 hr. potential gum, mg./l00 ml. 3,000 6. 8 '7. S Sulfur, wt. percent... O. 090 0 022 0.040 Nitrogen, p.p.m

Composition vol. percent: Paratlins 10. 2 16. 8 12. 8 11.3 9.4 14. 5 5. 3 0.0 O. 0 14. O 7. 2 ll. 1. 8. 2 15. 6 10. G 51.0 51. 0 51. 0 Sulfur, wt. percent 0.090 0. 022 0.040 Percent desulfurization 76 56 EXAMPLE 2 In another run, the charge treated was a dripolene naphtha boiling over the range of 117 to 312 F., having an API gravity of 45.4, and containing 4.5 p.p.m. nitrogen.

Hydrogenation was carried out over sulfided cobalt molybdate catalyst under the following conditions:

Temp, "F. 600 Pressure, p.s.i.g 600 Space rate, LHSV 2 H /oil ratio, mol/mol 4 The results of hydrogenation are shown in the following table.

TABLE 2 Charge Product Bromine N0 G7. 0 0. 8 Paratfins, vol. percent 14 2 28. 7 Olefins, vol. percent Mono-olefins Di-olefins-.. Cyclo-olefins N aphthenes, vol. percent... Aromatics, vol. percent" Styrene, vol. percent Sulfur, wt. percent chg.

Percent desulfurization... Existent gum, nag/100 ml EXAMPLE 3 TABLE 3.WEI GHT PE RCENT Hydrogenated Charge Product Paraifins Acyclic nomo-olefins 0. 09 Acyclic di-olefins 0.05

Di-0lefins.. Dicyclo-diolefi Styrene Vinyl toluene. Aromatics The total effluent from the pre-hydrogenation reactor was passed to the reactors containing chrome-alumina catalyst to effect hydrocracking of non-aromatics and dealkylation of alkylaromatics.

Operating conditions at the inlet of the cabalt-molybdate reactor were as follows:

Pressure, p.s.i.g. 800 H /hydrocarbon mol ratio 6 The fresh feed contained about 1.5 to 3 parts per million sulfur. In order to maintain the catalyst in fully sulfided state, there was added to the pyrolysis condensate feed stock prior to vaporization, 1000 parts per million of dodecyl mercaptan. The reaction effluent was collected during a 58 day run and processed over chrome-alumina catalyst as described below. The composition of the effluent from the cobalt molybdate reactor is shown in Table 3 without accounting for the hydrogen present. The chromealumina reactors (two in series) were operated at the described nominal conditions for two separate runs in the first run using the hydrogenation effluent alone and in the other run with the addition of paraffins and naphthenes to determine the effect of lowered aromatics concentration. The results are shown in Table 4 below.

In run C, the charge entered the first reactor of the two-reactor chrome-alumina system at 1092 F. and was discharged at 1174 F., cooled to 1095 F. at which temperature the total effluent entered the second reactor, leaving the said reactor at 1176 F. In run D, the charge entered the first reactor at 1090 F. leaving at 1179" E, was cooled to 1100 F., at which temperature it was ad mitted to the second reactor, leaving at 1175 F.

TABLE 4.RUN CONDITIONS (CHROME-ALUMINA CATALYST) Run C (without Run D (with added hydroadded hydro carbon) carbon) Temp. avg, F 1,135 1,129 Pressure, p.s.i.a.. 815 815 SV 0. 88 0. 04 Res. time rec 40 35. 5 Fresh I-Iz/arom m./m... 2. 56 4. ll Recyc. gas/arom., m./m 5. 41 5. 83

Run 0 Run D Net feed, Net prod, Net feed, Net prod., mol mol mol mol C omposition percent percent percent percent H2 68. 85 50. 73 75. 82 50. 02 C1 to O paraf. and

01 f1 0.00 21.17 0.00 20. 74 0.01 0. 20 0.01 0. 41 0. 22 0. O. 14 0. 19 0. 34 0.01 0. 70 0. 03 2. 31 0. 20 1. 0. 51 0.62 0.01 0. 0.02 Cyclohexane 0. 22 0. O0 0. 83 0. 00 Methyl cyclohexane.. 0. 24 0.00 0.89 0. 00 Benzene 15. 13 23. 54 10. 35 13. 89 Toluene 8. 05 3. 41 5. 34 4. 42 Ethyl benzene 1. 09 0.18 0. 78 0. 29 Xylenes 2. 29 0.14 1. 72 0.33 C aromatics. 0. 25 0.00 0. 22 0. 02 Naphthalene 0. 01 0. 09 0. O0 0. 07 Diphenyl-.. 0. 00 0. 11 O. 00 0. 03 C and C3 non-cy r O. 26 0.00 0.78 0.00 Unidentified 0. 09 0. 01 O. 07 0. 01

Product quality Freeze point, C 5. 54 5. 40 Thiophenes, p.p.m.. 0. 10 0. 10 Bromine index 0.20 2. 80 Liq. prod. residue, wt. percent. 0.07 0.17 Benzene purity, wt. percent 00. 988 99. 954

The process of the invention can be effectively utilized for conversion of any impure aromatic hydrocarbon stream containing 0,; to C hydrocarbons to extremely pure product benzene in a single integrated plant. Feed vaporization is effected continuously without fouling of heat exchangers and all sections of the plant can be operated continuously for periods of three months or more without necessitating regeneration or replacement of catalyst. Moreover, the catalyst in each section can be restored to activity by simple regeneration in dilute air.

What is claimed is: 1. A process for the production of benzene containing less than 1 p.p.m. of sulfur from a pyrolysis naphtha containing polymerizable unsaturates, which comprises: adding sulfur as needed to said naphtha to maintain from a minimum of 0.005 to 0.5% by weight of sulfur therein, whereby said naphtha can be vaporized with minimum polymerization -of unsaturates;

vaporizing said naphtha by direct sensible heat transfer;

contacting said vaporized pyrolysis naphtha charge stock and a hydrogen-containing gas in a first reaction zone with sulfided cobalt molybdate containing a total of about 8 to 20% by weight of the oxides of molybdenum and cobalt in the approximate weight ratio of about 3.5/1 to /1 (MoO /CoO) under conditions eifecting hydrogenation of readily polymerizable unsaturates;

contacting the efiluent from the first reaction zone with chrome-alumina catalyst which contains 15 to 25% Cr O incorporated in an alumina base derived from dehydration of hydrous alumina containing at least 50% beta trihydrate in a second reaction zone under hydrocracking conditions, said reaction conditions being such that substantially all of the sulfur in the naphtha feed is converted to H 5;

separating the efiiuent from the second reaction zone into a liquid phase and a vapor phase; and

recovering product benzene free of H 8 from said liquid phase.

2. The process as defined in claim 1, characterized in that the vaporized mixture of charge stock and hydrogen-containing gas is contacted with the catalyst in the first reaction zone at an inlet temperature of between about 400 and about 650 F., and the outlet temperature in the second reaction zone is between 1100 and 1230 F.

3. The process as defined in claim 1 characterized in that the pyrolysis naphtha is prefractionated to obtain a fraction boiling in the range of about 150 to 300 F., and this fraction is vaporized with hydrogen-containing gas to provide said vaporized mixture.

4. The process as defined in claim 3 wherein sulfun compound is added to the fraction subjected to vaporization, to bring the sulfur content of the fraction to at least 0.01% by weight of the hydrocarbons therein.

5. The process as defined in claim 1, characterized in that product benzene is recovered from said liquid phase by stabilizing said liquid phase through the removal of hydrogen, H 5 and light hydrocarbons, clay treating the resulting sulfur-free stabilized material to remove trace olefins and then fractionating the resulting clay treated material to obtain product benzene.

6. The process as defined in claim 1 characterized in that the second reaction zone comprises at least two reactors with means for cooling the intermediate reaction products.

7. The process as defined in claim 6 characterized in that at least part of said vapor phase separated from the second reaction zone efiiuent is employed in cooling said intermediate reaction products.

8. The process as defined in claim 1 characterized in that at least part of said vapor phasematerial separated from the second reaction zone efiluent is employed in the vaporization of the charge stock.

9. A process for the production of substantially sulfurfree high purity benzene from a dripolene charge stock, which comprises:

fractionating said charge stock to obtain a heat fraction boiling in the range of about 150 to 300 F.; vaporizing said heat fraction with hot hydrogen in the presence of sulfur compounds in quantity sufiicient to maintain a minimum of from 0.005% to 0.5% by weight of sulfur therein and to inhibit polymerization of unsaturates in said fraction;

passing the vaporization effluent over sulfided cobalt molybdate catalyst to effect hydrogenation of the readily polymerizable unsaturates contained therein, said hydrogenation being effected at a reactor inlet temperature of 400650 F. and at a liquid hourly space velocity of 0.1 to 5.0; passing the hydrogenated effluent with residual hydrogen therein over chrome-alumina catalyst containing 15 to 25 Cr O incorporated in an alumina base derived from dehydration of hydrous alumina c0ntaining at least 50% beta trihydrate under conditions favoring hydrocracking of non-aromatics without cleavage of aromatic hydrocarbon rings, said conditions including temperatures of 1100 to 1200 F.

and pressures of 800 to 1200 p.s.i.g., the sulfur in the effluent product from said reactor being present in the form of H 8; and

recovering benzene free of H 8 from the hydrocracking efiluent product.

References Cited UNITED STATES PATENTS DELBERT E. GANTZ, Primary Examiner A. RIMENS, Assistant Examiner U.S. Cl. X.R. 20889; 260672 'zg g UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 3 +87, 120 Dated December 30, 1969 Inventor(s) Robert G. Craig et a1 It is certified that error appears in the above-identified patent and that said Letters Patent are hereby corrected as shown below:

I Column 2, line 39, 'here" should read --there--.

Column line 46, "advantages" should be singular.

Column 5, line 53, "vaporizer 1" should read --vap0rizer l6--.

Column 7, line 9, table 1, "86.4" should read --68. l--.

Column 7, line 63, table 3, "nomo-oleiins should read --mono-olefins--.

Column 7, line 68, the words vinyl toluene" and "aromatics are bracketed together instead of the words 'styrene" and "vinyl toluene".

Column 10, lines 16 and 18, "heat" should read -heart--.

SIGNED AN scum JUN 9 1970 @EAL) Atteat:

y wmnrm E. sumnmm. JR. Edward M. Fletcher, Ir. g issione'r of Patents Attesting Officer 

